Refining of synthetic hydrocarbon mixtures



Jan. 8, 1952 R. v. sHANKLAND ETAL REFINING OF SYNTHETIC HYDROCARBONMIXTURES Filed Nov. 8, 1947 Patented Jan. 8, 1952 REFINING F SYNTHETICHYnRocARBoN MIXTUREs Rodney V. Shankland, Chicago, Ill., Stanley E.Shields, Whiting, Ind., and Ernest W. Thiele, Chicago, Ill., assignorsto Standard Oil Company, Chicago, Ill., a corporation of IndianaApplication November 8, 1947, Serial No. 784,880

Claims. 1

This invention relates to the refining of synthetic hydrocarbonmixtures, such for example as are obtained by synthesis from hydrogenand carbon monoxide. The invention pertains more particularly to animproved method and means for obtaining high quality distillate fueloils as well as high quality gasoline.

When hydrocarbons are synthesized by reaction of carbon monoxide andhydrogen over an iron catalyst, the gasoline fraction of the productliquid has a relatively low octane number and a bad odor and it isextremely unstable toward oxygen, i. e..toward gum formation anddiscoloration. The fraction boiling above gasoline is likewisecharacterized by bad odor and instabil- .ity toward oxygen so that it isunsuitable for use as distillate fuels such as kerosene, burner oils andfurnace oils. The obtaining of maximum yields of distillate fuel oils isa matter of outstanding importance; such products were heretofore amplysupplied as by-products from ordinary refining operations but theincrease in demand therefor coupled with the decrease in crudepetroleums of suitable quality for supplying such demand has presentedthe petroleum industry with a serious problem which the presentinvention will help to alleviate. An object of this invention is toprovide a method and means for treating the synthesis product liquid forproducing maximum quantities of high quality distillate fuels andgasoline by a reaction which is primarily reforming and deoxygenation asdistinguished from cracking..

A further object is to provide an improved method for treating synthesishydrocarbon product mixtures which Will minimize degradation thereof tocarbon and light hydrocarbon gases. A further object is to obtain fromsynthesis hydrocarbon mixtures maximum yields of gasoline and distillatefuels at minimum refining costs and to provide a renning system whichcan be built and operated at minimum expense.

A further object of the invention is to provide an improved treatingmethod for deodorizing synthetic hydrocarbon products of the gasolineand distillate fuel boiling range and for making said products stableagainst discoloration and gum formation. Another object of the inventionis to provide optimum operating conditions for effecting such treatmentwith particular catalysts. A further object is to provide an improvedcorrelation of temperature, catalyst and severity of treatment wherebymaximum yields of high quality distillate fuel oil and gasoline areobtainable from synthesis hydrocarbon mix- 2 tures with minimumdegradation to coke and light hydrocarbon gases. A further object is toprovide improved methods and means for remov.- ing combined oxygen fromsynthesis hydrocarbon products. Other objects will be apparent as thedetailed description of the invention proceeds.

In a preferred method of practicing the invention the synthesis productliquid is heated to effect vaporization of substantially all componentsexcept those which might form carbonaceous deposits 'in heating tubes,the vaporized portion is separated from unvaporized liquid, and thevaporized portion is then superheated. Unvaporized liquid may becommingled with the superheated vapors when said vapors are contactedwith conversion catalyst. The catalyst is preferably a synthetic silicaalumina gel which contains about 10% to 25% of alumina` and which hasbeen dried and heated to high temperature to produce hard, porousparticles characterized by a gel structure. The catalyst is of smallparticle size chiefly in the range of 2 to microns and itr is employedin fluidized dense phase condition in a reaction and regenerationYsystem similar to that conventionally employed in so-called fluidcatalytic cracking systems.

The severity of treatment in our process is materially lower than thatemployed for .catalytic cracking. The reaction temperature of about 700F. is much lower than ordinary cracking temperature. The intensity oftreatment depends not only upon temperature but upon the activity of thecatalyst, the catalyst-to-oil weight ratio (C/O) and the weight spacevelocity or amount by weight of oil charged per hour per amount byweight of catalyst in the reactor (Wo/hr/Wc) After catalyst has been onstream for some time its activity is considerably less than that offreshly prepared catalyst and its relative weight activity (A) may bedefined as the number of parts by weight of fresh catalyst which wouldcorrespond in effectiveness to 100 parts by weight of the catalyst beingevaluated. Equilibrium catalyst in a fiuidized system may have arelative weight activity in the range of about 10 to 40, e. g. about 15to 20. At such catalyst activity the severity factor may be expressedas:

and the severity factor as determined by this formula should be in theapproximate range of severity factors greater than 2.0 result in im-Ypaired quality and an undesirable amount of cracking, loss to coke andgas formation.V

The reactor is operated at a dilute phase pressure of about 12 poundsper square inch gauge,

' the pressure in the lower part of the reactor' being somewhat highersince the catalyst therein is maintained in turbulent condition at adensity of about to 35 pounds per cubic foot by passing a gasiformcharge streamV upwardly therein at about 1.2 to'1.5 feet per second(measured in the dilute phase).

Catalyst regeneration' is eiected at about 1000 to 1050 F. at adilutephasepressure of approximately 1'0 pounds per square inch gauge.The catalystV in the lower part' ofthe regenerator is nraintained inturbulent condition at a density of about 25 to 35 pounds per cubic footby passing regeneration gas upwardly therein at about 1.3 to 1.6 ieetper second (measured in the dilute phase). The catalyst is stripped withsteam before it enters the-regeneratcr from the reactor.

i The products from the reactor enter the lower scrubbing section ofafractionatorrwherein residual catalyst particles are scrubbed out andsettled for return to the stream entering the reactor. Naphtha, heateroil and gasoil streams are separately recovered. In the absence ofpreliminary adsorption` of oxygen compounds from the synthetichydrocarbon mixture with silica gel or thelike, the heater oil and/orgas oil Y streams should either be recycled to theI reactor withincoming synthesis hydrocarbon product or maybe recycled in a blockedout operating in the absence ofiV synthesis hydrocarbon product. Theseparticular fractions require more drastic treatment than the gasolinefraction for deodory about 600to 700 F. and a pressure in theV rangeV ofabout 200 to 450 pounds per square inch, the total eilluent synthesisstream being fractionated to remove normally gaseous components andoxygen compounds that are readily removable by feasible fractionation,extraction or other separation means. The remaining synthesis hy*-dr'ocarbon mixture consists essentially of hydrocarbons having about 5to 20 carbonatoms or more per molecule with combined oxygen in amountsof .5 to 5% or more.

The treatment of fractions containing less than 5 Vcarbon atoms permolecule or even the C5 fraction is usually unnecessary because oxygencompounds can readilybe removed from such fractions by simple and wellknown extracf tion Y or v fractionation procedures. products havingappreciably more than 20 car- Any waxy bon atoms per molecule may beremoved by distillation or other known means. The inspection of arepresentative synthesis hydrocarbon mixture is as follows:

API gravity Y 55.2 Combined oxygen 2.12 weight Initial boiling point 128F. 10% over at 168 F.

30% over at 218 F.

% over at 270 F.

% over at 339 F.

% over at 474 F.

VEnd point 650 F.

The combined oxygen is usually rather uniformlydistributed throughoutthe entire boiling range. This synthesis hydrocarbon mixture usuallycontains about 78 to S2 volume percent of hydrocarbons in the gasolineboiling range, a representative inspection of the gasoline fractionbeing approximately as follows:

API gravity 61.1v Odor bad Color 3.5 NPA ASTM gurn 111.4 mg. Inductionperiod (with and Without inhibitor) 40 min.

Aniline point, oF 74 Oxygen content 2.16 wt. Clear ASTM motor octanenumber 58.8

Motor octane number with 1 co. of

lead tetraethyl 66.4 Motor octane number with 3 cc. oi

lead tetraethyl v '74.1 CFR research octanernumber clear 63.8 CFRresearch octane number-i-l cc.

of lead. tetraethyl 70.4 CFR research octane number-F3 cc.

of lead tetraethyl V83.7 Initial boiling point 122 F.` l 10% over at 166F. 30% over at 196 F. 50% over at 228 F. '70% over at 266 F. 90% over at328?. End point 395 F.

The fraction of the charge in the heater oil range is characterized by abad odor and a marked tendency toward discoloration and gum formation onstorage. Y

About 4500 barrels per day of total synthesis hydrocarbon mixture ofthis general type is chargedrfrom source l@ by pump l I through heatexchanger i2 and line. i3 to preheater llwhere-Y in it is heated V to asuicient temperature for vaporizing all but the highest boilingfractions of the charge, e. g. to a temperature in the range of 500 to600(F. or about 550 F. The heated charge is then introduced intoseparator l5 from which the vapors may be passed by line 'it throughcoils in superheater furnace il and thence passed by line i3 to transferline id or picking up hot catalyst discharged from the base of standpipe20 in amounts controlled by valves A by-pass may be provided around heatexchanger !2 and around superheater Il `for at least a part of thestream normally passing through lines i3 and I8 in order to obtain de'-sired temperature control. The -unvaporized liquid from the base Vofseparator lois withl drawn through line 22 and it may be in whole or inpart introduced by line 23 to transfer line i9 or introduced by line 24to the base of the `highest boiling components product 'fractlonatingtower. The vapors passed through superheater I1 may be heated to atemperature of' the order of 650 to 700 F. the precise amount of preheatdepending upon the amount of unvaporized charge introduced through line23 to transfer line I9 and the amount of heat available as sensible heatin the hot regenerated catalyst. By eliminating the from materialspassing through the superheater coke deposition in the superheater`coils is substantially avoided and coke formation in the reactor issubstantially reduced.

The catalyst-to-oil weight ratio of materials passing through transferline I9 to the reactor is preferably in the range of about 2:1 to 5:1and it may be varied within or even outside of this range to maintainthe desired heat balance, i. e. to maintain the temperature in thereactor at about 700 F. The hot charge and catalyst mixture passesthrough transfer line I9 at a velocity of about 25 to 30 feet per secondto the base of reactor 25 into which it is distributed by grid plate 26designed to give a pressure drop of about l pound per square inch and toeffect uniform distribution of the incoming mixture throughout thecross-sectional area of the reactor. The reactor itself may be about 40feet or 50 feet high and about 11 feet in diameter so that in recycleoperation the vertical gas velocity in the upper portion thereof will bein the range of 1.2 to 1.5 or about 1.35 feet per second. Under suchoperating conditions the catalyst density will be in the range of about25 to 35 pounds per cubic foot in the lower part of the reactor. Thepressure in the dilute phase above the dense phase level is of the orderof about 10 to 12 pounds per square inch gauge and a dilute phase ordisengaging space of atnleast about feet should be maintained above thedense phase level. For average equilibrium catalyst activity the weightspace velocity in the reactor should usually be in the range of about 2to l0 pounds of oil charged per hour per pound of catalyst in thereactor at any instant.

The reaction products pass from the upper part of the dilute phase tocyclone separator 21 from which separated catalyst particles arereturned by dip leg 28 to a point below the dense phase level in thereactor. The product stream then passes by line 29 to the lower part offractionator tower 30 wherein high boiling product components arecondensed and solids are scrubbed out of the product stream. Condensateand solids are withdrawn through line 3| to settling chamber 32. Settledsolids may be returned from the base of the settler by line 33, pump 34and line 35 to transfer line I3 along with a small portion ofthe chargestream introduced through line 38. Another portion of the liquid fromthe base of tower 30 is withdrawn through line 3'! and forced by pump 38through exchanger I2 and then returned by line 39 as scrubbing liquidfor the lower part of tower 30. For further temperature control a partof this recycled stream may be passed through auxiliary temperaturecontrol means 40. Solids settling chamber 32 is connected by vent line4l to a point in tower 30 above the inlet of line 29. The heavy oilseparated from solids in settler 32 may be Withdrawn by pump 42 andeither discharged from'the system by line 43 or returned throughlines 44and 45 to inlet charge line I3.

A heavy gas oil side stream is withdrawn from tower 30 through line 46to stripper 41 into which steam is introduced through line 48, thestripping steam and overhead fraction being returned to the towerthrough line 49 and the heavy gas oil fraction being withdrawn by pump50 and line 5I or returned by lines 52 and 45 to inlet charge stream inline I3.

A heater oil fraction boiling in the range of about 350 to 600 F. iswithdrawn from tower 30 to line 53 to side stream stripper 54 wherein itis stripped with steam introduced through line 55 the steam and overheadproducts being returned to the tower through line 5E and the heater oilbeing withdrawn by pump 51. All or a part of the heater oil is withdrawnfrom the system through line 58 although a substantial part of theheater oil may be returned by line 45 to the charge stream in line I3.

A heavy naphtha fraction is withdrawn from tower 30 through line 59 tostripper 60 into which steam is introduced through line 6I and fromwhich overhead is returned to the tower through line 62. The heavynaphtha is withdrawn from the system through line 63.

Light naphtha together with uncondensed gases leave the top of tower 30through line 94 to cooler-condenser 65 and is thence introduced intoAseparator' 66 from which condensed water is withdrawn through line 61.Uncondensed gases leave separator 66 through line 68 and aresubseduently processed to recover the desired components thereof.Condensate from the base of separator 6E is withdrawn through line 65 apart of it being returned by pump 'I0 for use as reflux and theremainder being withdrawn through line 10a for stabilization.

Referring to the catalyst system generally, a substantially constantcatalyst inventory is maintained therein and relatively spent catalystmay be withdrawn from the system and fresh catalyst added thereto atsuch a rate as to maintain a predetermined weight activity which may,for example; be about 20. Catalyst is removed from the reactor atsubstantially the same rate as it is introduced thereto from theregenerator, the catalyst from the reactor being withdrawn directly fromthe dense phase at a point below the upper level thereof. When thestripper is at the base of the reactor and in the same chamber, thecatalyst may be withdrawn through annular space 'II to stripping sectionI2 into which steam is introduced through line 'I3 at the rate of about6 pounds per thousand pounds of catalyst passing through the stripper.If desired, an external stripper may be employed in which case densephase catalyst may be transferred laterally thereto through a valveconduit and the stripped products may be returned by an upper conduit tothe reactor. The stripping steam replaces the hydrocarbon suspension gasand even when the stripper is in the base of the reactor very little ifany steam passes upwardly into the dense catalyst phase in the reactor.We have found that it is detrimental to introduce any substantial amountof steam into the dense phase portion of the reaction zone as will behereinafter pointed out in more detail.

The stripped catalyst is downwardly withdrawn through standpipe 'I5 andintroduced in amounts regulated by valve 'l5 into transfer line 'Ilwhich leads to regenerator 18. From blower 'I9 a portion of the air maypass directly through lines and 8l to the base of the regenerator butsuilicient air is introduced through line |32, vessel B3 and line 84 forcarrying catalyst hack to the regenerator at a velocity'of about 25 to30 7 feetper second in transfer line AFL During the starting .upprocedure a .portion of the fair may .support combustionrin chamber 83of iuel gas in- :troduced Vthrough Aline 85 the -hotcombustion productsserving to .bring the catalyst up to desired temperature. If the amountof `carbon .deposited on 'the' catalyst in the combustion step is 'notsuiicient, when yburned in the regenerator, :to maintain desi-redregenerated catalyst temperature, torch oil may be. directly introducedvinto the regenerator through line 86.

The `regenerator may be vof substantially the 'samegeneral design asrthereactor but somewhat AGarger and in this case may be a vessel of about140"to=50 feet linheight and about 13'1/2 feet in diameter (makingallowance `for maximum coke to be handled). The regenerator ispreferably opera'tedfata temperature of about 1000 to 1050 F. and'atdilute phase pressure of about 10 to 12 pounds per square inch gauge.The incoming spent catalyst stream vis uniformly distributed across thelentire cross-sectional area by a grid 31 which may be designed for apressure drop off about 41 pound per vsquare inch. The regeneratorshould be designed for a dilute phase upward gas velocity intheV rangeof about 1.3 to 1.6, e. g. about 1.45 feet per second. Here again thedense phase-catalyst inthe lower part of the regenerator should have adensity of about 25 to 35 poundsper cubic foot, the bed depth may beabout feet and at least about 15 feet should be provided-abovethe densephase level for disengagement' of carry-over catalyst particles.Pri-mary,- secondary andtertiary cyclone separators88, 89 and 90 removethe bulk `of the remaini-ng-entrained catalyst particles from the fluegas and these particles are returned to the dense phase by dip legswhich extend below the dense phase level. Flue gas is discharged fromline 9! through a valve which controls the pressure in the top of theregenerator and, in order to protect suc-h valve means,l spray water isintroducedl through line 92 to cool the exit gases to a temperatureofapproximately 700 F.

Regenerated catalyst may be withdrawn di rectly from the dense phase inthe regenerator into the top of standpipe 20. It should be understood ofcourse that standpipes and 15 are both provided with conventionalaeration means at apoint immediately above valves 2l and 'I6 vand atother points along the standpipe lengths. Make-up catalysts may beintroduced into the system through line 93, emergency spray Water may be'introduced into the upper part of the regenerator throughk 94 andemergency steam may also be introduced through line 95 and transfer line@l at the base of fractionator 30.

YAs illustrative of results obtainable b-y our process, a charge, ashereinabove described, was treated with a iiuidized silica aluminacatalyst hai/ing a. particle size chiefly within the range of about '1to '100 microns and synthetically preparedA to have agel structure andto have an alumina content in the range of about 10 to 25%. The charge'was treated at about 700 F. with a weight space velocity ofV about 3.9pounds of charge per hour per pound of catalyst in the reactor and witha catalyst-to-oil Weight ratio of about 2.5:1 at a pressure .of about10.7 pounds per square inch, the regeneration of the catalyst beingeiected at about 1000 F. The catalyst had a relative weight activity ofabout 20 to 30 and the severity factor was about 0.33. On a ieasigaeoonce-through product output `basis therewa'sfobtained:

The C4-400 F.7.gasoline fraction had the following properties:

API gravity 61.5

Odor O. K. Color 11Saybolt Aniline point, F v97 ASTM -gum 22mg.

Induction period without inhibitor--. 48o min.

Induction with .003% du Pont No. 22 '705 min. ASTMfmotor octane No.clear 76.4 ASTM motor octane No. 1 cc. lead tetraethyl 81.6

ASTM motor octane No. 3 ccflead tetraethyl .84.3. CFR-R octane No. clear85.3 CFR-R 4octane No. -il cc. lead tetraethyl V 91.4 CFR-R octane No. 3cc. lead tetra- .ethyl `95.3 Reid vapor pressure 4 lbs. Y Initialdistillation temperature 120 F.

10% over at 16.6 F. 30% over at 204 F. 50% over at 239 F. '70% over at283 F. 90% over at 350 F. End point approximately 390 F.

The heater oil, which constituted 9.3% of the total product, 'had thefollowing inspection:

API gravity 35.4 Odor Fair Color 1.5 NPA Initial boiling point 436 F.10% over at 448 F. v30% over rat 458 F. 50% over at 467 F. 70% over at480 F. 90% over at 506 F. End point 528 F.

While the above inspection shows lthat a remarkable improvement waseffected in the heater oil `over theV corresponding fraction in theinitial charge, best results are obtainable if a substantial portion ofthe heater oil-fraction, as well as the heavy gas oil fraction, isrecycled to the conversion zone or separately retreated with thecatalystat substantially the same temperature. Thus when about two partsof product boiling above 400 F. is recycled for each three parts offresh charge introduced, the dry gas yield may be increased byabout 3%,the total C4 hydrocarbons may be increased' by about the same amount,the gasoline fraction' may be increased by about l or 2% and the heateroil fractiondecreased by about 1 or 2%, slightly more water beingproduced inthe recycle operation and the coke being increased by about.8% based on total charge.

Such a recycle operation markedly improves the heater oil fraction withregard to odor, color and stabl'ityragainst gum formation. In therecycle operation we may recycle all of the heavy gas oil and all butthe net production .of the heater oil. For best results the recycleratio of total' feed heater oil was good after 500 hours.

15,0 fresh feed should be in the range of 1.2:1 to 3:1; based oncomponents in the fresh feed boiling above 400 F. the ratio should be inthe range of 2:1 to 11:1.

Instead of employing recycle` operation the heater oil and/or heavy gasoil may be separately charged to the reactor at about 700 F. undersubstantially the same conditions as hereinabove set forth. In fact thegasoline fraction and the higher boiling fraction of the originalsynthesis mixture may be separately treated in which case the gasolinefraction may be subjected to somewhat higher temperatures andcorrespondingly higher space velocities and the heater oil and gas oilfractions may be contacted at temperatures of approximately 700 F. atsomewhat lower space velocities, e. g. of the order of 0.1 to 3 poundsof oil per hour per pound of catayst in the reactor at any instant. Thusa'fraction of the original synthesis hydrocarbon mixture having an APIgravity of 35.4 and a distillation range of about 400 to 675 F. wastreated with silica alumina catalyst at a temperature of 700 F. with aweight space velocity of .22 pound of oil per hour per pound of catalystin the reactor at a pressure of 15 pounds per square inch to give a 92%volume yield of a product of whichA the kerosene and heater oilfractions, on standing 100 hours in the daylight at room temperature,had a saybolt color shade of plus 20 and plus 11 respectively. The odorof the kerosene and The color stability of the heater oil was evengreater than that of kerosene produced from crude petroleum.

While a fluidized treating system has been hereinabove described inconsiderable detail, it should be understood that the treating may beeffected in either fixed bed or moving bed operations although suchoperations (particularly xed bed) are not as advantageous, economical ordesirable as the fluid catalyst system. When fixed bed or moving bedoperations are employed, the on-stream` period (or catalyst holdingperiod in movingbed) should be less than 2hours`and preferably fromabout 1 minute to about l hour, the space velocity in this case beingwithin a range of about .1 to 2.5 pounds of oil charged per hour perpound of catalyst in the reactor at any instant. The use of shorton-stream reaction periods between catalyst regenerations has aremarkable effect in increasing the rate of conversion and productquality and in decreasing the losses to'dry gas and coke. Since it isdesired that cracking of hydrocarbons, particularly in the distillatefuel boiling range, should be minimized and that the conversion shouldbe chiey reforming and deoxygenation, the very short catalyst holdingtime (or on-stream period in fixed bed operations) is a matter ofconsiderable importance.

The reaction temperature of 700 degrees is predicated on the use ofaverage equilibrium catalyst originally prepared as a synthetic,gel-type silicaalumina catalyst containing about to 25% of alumina(altho the alumina content may be more or less than that stated as thepreferred range). An example of the preparation of such a catalyst is asfollows: dissolve about 50 to 75 kg. of sodium aluminate in about 500liters of water. If the sodium aluminate is of suiiieient purity thesolution should be substantially complete. If a technical grade ofsodium aluminate is employed the solution should be filtered to removeinsoluble materials. Dissolve 700 liters of vsodium silicate .(Waterglass) in 2100 liters of water. Pour the sodium aluminate. into thewaterglass solution at ordinary room temperature `with stirring to avoid anyprecipitation and to obtain a sol of uniform consistency. Permit the solto. set for several hours orlonger to form a gel. Wash the gelthoroughly with distilled water and then with a weak acid or aconcentrated salt solution, e. g., an 8% ammonium chloride solution, forremoving `sodium ions as completely as possible. Then thoroughly washthe gel with water and slowly dry the gel to such degree that it can bebroken int'o small lumps. Continue the drying of the lumpsl of brokengel at higher temperatures and finally heat the lumps of gel to atemperature of about 850 to 1150 F., preferably about 1000A F. Theresulting product will consist of hard glassy appearing (i.` e. `not`dull' or chalky) `particles which are remarkably resistant toabrasionand which are characterized by a high degree of porosity. Theresulting catalyst is a dense form of silica and alumina which is inintimate physical admixture, which is highly porous and which hasremarkable activity. Synthetic silica alumina catalysts of similaractivity can be prepared by other methods well known to those skilled inthe art and the invention is applicable to the use of any suchcatalysts.

Considerable improvement in gasoline and heater oil properties may beobtained by the use oi' activated montmorillcnite clay (Super Filtrol)or by the use of activated alumina catalysts pre` pared from aluminagels or bauxite. Such activated alumina catalysts however are notequivalent to the synthetic silica alumina `catalysts hereinabovedescribed and they require somewhat dirferent operating conditions. Whenactivated alumina catalysts are employed the treating temperature shouldbe about 100 degrees higher and the contact time should be somewhatlonger, i. e. the weight space velocity should be somewhat lo'wer andthe severity factor corre. spondngly increased. Activated aluminacatalysts tend to form more carbon or coke and to` result in somewhatlower product quality. The use of fullers earth (Gray process) Vat 500F., 200 p. s. i. g. and a weight space velocity of 1 pound of oilcharged per hour per pound of fullers earth in the treating zone gavesome improvement in stability but no appreciable irnprovement in odorand insufficient improvement in octane number; the failure of thistreating method to produce a satisfactory gasoline is` striking evidenceofthe fact that the odor andY instability of the synthetic hydrocarbonmixture are caused by very different typesof compounds from those whichare responsible for instability and/or poor color in ordinary petroleumrefining processes.

When we refer to a temperature of about 700 we mean 700 F. plus or minus75 F. Similarly, in the case of bauxite,` a temperature of about 800means a temperature of 800 F. plus` or. minus F. Preferably, thetemperature should be in the range of plus or minus 50 F. and for bestresults should be in the range of plus or minus 25 F. Highertemperatures produce too much cracking, coke deposition and gasformation and result in products of lower quality. Lower temperatures donot accomplish the desired product improvement at reasonable severityfactors.

Addition of steam to the reactor at higher temperatures tends tosomewhat reduce the coke deposition but in this process steam has adeleterious effect on the stability of the gasoline amount of waxybottoms is not included).

atenten:

Iii obtained and its use also impairs the odorlof the gasolineproduced;v the use of steam seems to destroy the effectiveness of theprocess toward the removal of oxygen .in the synthetic hydrocarbonmixture. Although small amounts of steam may be tolerated in thereactor, e. g. the amounts that might come from the stripping zone, Weprefer to avoid Vthe introduction of steam into the dense phase catalystyportion of the reactor.

As hereinabove pointed out, the severity of treating is of greatimportance. The severity Vshould be such as to avoid cracking as apredominant reaction and to effect instead chiefly isomerization Yandoxygen removal. The term isoforming has been applied to this treatingstepbecause of the high liquid yields, octane number improvement andrelatively small amount of cracking; it should be pointed out, however,.that this process is markedly different from the process described, forexample in U. S. Letters Patents 2,326,705 and 2,410,908, in ythat theproblem in this case is that of removing or altering very diierent typesof compounds,Y the temperature employed is of a different order ofmagnitude, the optimum severity factor is likewise considerablydiiferent and the use of steam, which was beneficial in the priorprocess, appears to be detrimental in this process. In the priorisoforming process the charge to the treating step was fractionated toimprove components boiling above 400 to 450 F. While in this processadvantageous results are obtained by treating the entire synthetichydrocarbon mixture'except perhaps for a very small amount of heavy waxymaterials the amount of which is usually less than about In our processthe coke yield appears to be approximately the same under a given set ofoperating conditions regardless of Whether the feed stock is the gasolnefraction, the gas oil fraction or the total synthetic hydrocarbonmixture' (provided that the small Considerable economy is thereforeeffected by simultaneously treating the total gasoline plus gas oilfraction and then vrecycling lor retreating the heater oil fraction toobtainv the desired odor, color, stability against oxidation and gumformation and burning quality index. Our treating ofthe heater oilfraction does not improve its cetane number and for preparation ofdiesel fuel it is preferred to subject such fractions tohyolrogenation.V

As above pointed out, the severity of the treating must be maintainedwithin narrow limits, i. e. the severity factor, as above defined,should be inthe range of about .02 to 2.0 at about 700 F. with averageequilibrium catalyst. Such se- Verity factor sharply distinguishes thisprocess from the prior proposals to subject synthetic hydrocarbonmixtures to catalytic cracking. Thus U. S. 2,264,427 teaches thecontacting of synthesis hydrocarbon productmixtures and particularly thegas oil fractions thereof with conventi-onal cracking catalysts undercracking conditions with the object of converting the higher boilingcomponents to gasoline. In our process We seek to obtain deoxygenationWithout substantial cracking so that we may avoid the large losses tocoke and fixed gases, effect substantial savings in operating costs, andproduce a high quality heater oil as a separate fraction instead ofconverting such heater oil fraction to gasoline by destructive cracking.

l the synthesis hydrocarbon mixture 1s con- 2 tacted with an adsorbentsuch assilica gel-priorto our treating step ther severity of` treating're'-v quired for obtaining good. odor and stability isV somewhatdecreased but the indicated severityl of treatment may be required toobtain aproduct of desired octane number improvement. If: such a silicagel adsorption step. is employed for` minimizing the amount of oxygencompounds in vthe mixture undergoing treatment, a once-4 throughoperation may be all that is required to `provide a distillate fuel ofrequired odor and stability. While the use of such an adsorption stepmay be employed in our invention with advantageous results, theinvention is not lim-f.

ited thereto.

While our process is specifically designed' for the treating ofsynthesis hydrocarbon mixtures obtained from synthesis with lhydrogenandcarbon monoxide over iron catalyst, it is believed that the processmay also be applicable `to the treatment of synthetic hydrocarbonmixturesob-V tained by oxidation of hydrocarbons and/or by thedistillation of shale provided that the products in those gases containthe same types of objectionable components that are vcontained in thecharging stocks hereinabove described. The term synthetic hydrocarbonmixture asused herein is intended to mean a hydrocarbon mixturecontaining oxygen compounds which cannot be separated therefrom byconventional fractionation or extraction procedures, the amount ofcombined oxygen in said mixture usually being in ghe range of about .5%to 5% or more, e. g. about While our invention has been described inconsiderable detail with respect to specio examples thereof, it shouldbe understood that many modifications of apparatus andoperating-procedures and alternative operating conditions will beapparent to those skilled in the art from the above description.

We claim:

1. The method of treating a synthetic mixture of hydrocarbons andoxygenatedcompounds, which mixture boils Within the rangeY of about F.to about 700 F. and contains ohiey compoueuts in thoeasouue boilingrango. which mixture contains combined oxygen distributed throughout theentire boiling rango and present in an amount `in the. .rango of to 5weight por Cent and which mixture is oharaotoriaod by a had odor andextreme instability toward oxygon'wth a marked tendency towarddiscoloration and sum formation, which method oomprisos contacting, Saidmixture with a silica oatalyst of gol Structure containing an amountV ofalumina in the range oi about `10% to 25%V by weight, which issuiiciently spent so that its catalytic activity is only'one-tenth tofour-tenths the catalytic activity of fresh catalyst, at a temperaturein the range of 625 F, to 775 F. and with a, vseverity factor in therange of .02 to `2,0 sufficiently low to substantially avoid crackingand to prodllfo a.

product of. good odor and stability toward oxygen.. said severity factorbeing the catalyst-.to-oil weight ratio'divideu by the weight Spaovelocity taken to the lkrpowor,l

2. The method of claim l wherein the mixture.

d er a pressure in the .rango o f about 20.0 to 4,50'

pounds per square inch, the total eiuent synthesis stream beingfractionated to remove normally gaseous components, higher boilingcomponents and extractable oxygen compounds.

3. The method of claim 1 which includes the step of recontacting aportion of the product from a previous contacting step, which portion ishigher boiling than gasoline, said recontacting being under conditionsfor the production chieily of a heater oil of good color and stabilitytoward oxygen.

4. 'I'he method of claim 1 which includes the steps of employing aweight space velocity in the contacting step in the range of about 2 to10 pounds of oil introduced per hour per pound of catalyst in thecontacting zone and effecting said contacting in the absence of anysubstantial amount of added steam.

5. The method of claim 1 which includes the steps of superheating vaporsof said mixture prior to the contacting step and effecting saidcontacting by commingling superheated vapors with hot catalyst of smallparticle size and passing said vapors upwardly through a yfluidizeddense phase mass of said catalyst while employing a catalyst-to-oilweight ratio of materials introduced in the contacting zone in the rangeof about 2:1 to 5:1, a weight space velocity in the contacting zone inthe range of about 2 to 10 pounds of oil introduced per hour per poundof catalyst maintained in the contacting zone, and fractionatingproducts from the contacting zone to obtain a gasoline boiling rangefraction of improvedy octane number, good odor and stability towardoxygen and a burning oil boiling range fraction also characterized byimproved odor and stability toward oxygen,

R'ODNEY V. SHANKLAND.

STANLEY E. SHIELDS.

ERNEST W. THIELE.

REFERENCES CITED The following references are of record in the file ofthis patent:

UNITED STATES PATENTS Number Name Date 2,059,495 Smeykal Nov. 3, 19362,264,427 Asbury Dec. 2, 1941 2,360,463 Arveson Oct. 17, 1944 2,404,340Zimmerman 1 July 16, 1946 2,424,467 Johnson July 22, 1947 2,443,673Atwell June 22, 1948 2,476,788 White July 19, 1949 FOREIGN PATENTSNumber Country Date 112,274 Australia July 3, 1939 735,276 Germany May11, 1943

1. THE METHOD OF TREATING A SYNTHETIC MIXTURE OF HYDROCARBONS ANDOXYGENATED COMPOUNDS, WHICH MIXTURE BOILS WITHIN THE RANGE OF ABOUT 100*F. TO ABOUT 700* F. AND CONTAINS CHIEFLY COMPONENTS IN THE GASOLINEBOILING RANGE, WHICH MIXTURE CONTAINS COMBINED OXYGEN DISTRIBUTEDTHROUGHOUT THE ENTIRE BOILING RANGE AND PRESENT IN AN AMOUNT IN THERANGE OF .5 TO 5 WEIGHT PER CENT AND WHICH MIXTURE IS CHARACTERIZED BY ABAD ODOR AND EXTREME INSTABILITY TOWARD OXYGEN WITH A MARKED TENDENCYTOWARD DISCOLORATION AND GUM FORMATION, WHICH METHOD COMPRISESCONTACTING SAID MIXTURE WITH A SILCA ALUMINA CATALYST OF GEL STRUCTURECONTAINING AN AMOUNT OF ALUMINA IN THE RANGE OF ABOUT 10% TO 25% BYWEIGHT, WHICH IS SUFFICIENTLY SPENT SO THAT ITS CATALYTIC ACTIVITY ISONLY ONE-TENTH TO FOUR-TENTHS THE CATALYTIC ACTIVITY OF FRESH CATALYSTAT A TEMPERATURE IN THE RANGE OF 625* F. TO 775* F. AND WITHK A SEVERITYFACTOR IN THE RANGE OF .02 TO 2.0 SUFFICIENTLY LOW TO SUBSTANTIALLYAVOID CRACKING AND TO PRODUCE A PRODUCT OF GOOD ODOR AND STABILITYTOWARD OXYGEN, SAID SEVERITY FACTOR BEING THE CATALYST-TO-OIL WEIGHTRATIO DIVIDED BY THE WEIGHT SPACE VELOCITY TAKEN TO THE 1.5 POWER.